Process for upgrading petroleum naphthas



Nov. 17, 196.4l

J. W. SCOTT, JR., ETAL PRocEss FOR UPGRADING PETROLEUM NAPHTHAS Filed May 12. 1961 v g /l V REAcToR /9 GAS LIQUID SEPARATOR HYDROGEN A TORNEY? United States Patent O 3,1575@ PRCESS FR *UPGRADNG PETRGLEUM NAPHTHAS .iohn W. Scott, r., Ross, Salif., and Frank M. Hazard, fir.,

lim Tanura, Sandi Arabia, assignors to California Research Corporation, San Francisco, Calif., a corporation of Delaware Filed May l2, 195i, Ser. No. 112,742 6 Claims. (6i. 20S-d) This invention relates to a process for improving the octane number and other characteristics of petroleum distillates, especially middle distillate fractions, and it is particularly directed to a process wherein the distillate is rst fractioned into a number of streams, each of which is then passed with hydrogen at elevated temperatures and pressures over a catalyst having both cracking and hydrogenation activity. in the ensuing reaction, wherein substantial hydrogen is consumed, the distillate fraction of each feed stream is converted to a product boiling below the initial boiling point of said stream.

This application is a continuation-in-par-t of our application Serial No. 637,328, filed ianuary 30, 1957, now abandoned, which 'm turn was a continuation-in-part of our application Serial No. 478,784, filed December 30, 1954, now abandoned.

The present invention is based on the discovery that excellent yields of high quality products can be obtained from any one or more of a variety of petroleum distillate fractions boiling in the range from about 175 to 850 F. by a process involving the following interrelated steps. The distillate is first fractionated into at least two and preferably into three or more feed fractions, each having an appreciable boiling range. Thereafter, each of the fractions so obtained, along with at least 1500 scf. of hydrogen per barrel thereof, is passed through an isomerization-cracking Zone provided with a catalyst made up of one or more hydrogenating-dehydrogenating components disposed on a support having active cracking characteristics, said zone being maintained at pressures of at least 600 p.s.i.g., and at temperatures between about 400 and 900 F. The space rate utilized may be varied within relatively wide limits, eg., from 0.1 to 15. The reaction occurring under these conditions results in a substantial net consumption of hydrogen as the feed (including any recycled stock) is converted to the desired high quality product fractions. Said product fractions are recovered from the respective effluent streams from the isomerization-craching zone, or zones, as the portions boiling essentially below the initial boiling point of the particular naphtha feed fractions concerned and may therefore be termed synthetic products. Said synthetic products have a low content of normal paraflins, and a high content of isoparaflins yand ring compounds. They are valuable for low freeze point jet fuel purposes, and are reformable to high octane gasolines in good yield. While good conversion-s (often exceeding 50%) to said synthetic products can be obtained in single-pass operations, this conversion can be increased to values between 70 and 100% by recycling to the catalyst reaction zone the individual bottoms streams remaining after removal of synthetic products, said recycle being returned to the reaction zone either aloneor along with added quantities of the particular feed fraction concerned, added hydrogen being also supplied in either case.

As indicated above, there is a net consumption of hydrogen, normally amounting to from S00 to 3000 s.c.f. of hydrogen per barrel of feed converted in the iso-merization-cracking zone. in this connection, conversion is arbitrarily expressed in terms of the amount of feed (including recycled eiiluent bottoms) converted to product Patented Nov. 17., 1964 ice boiling below the initial boiling point of said feed stocks. This hydrogen deficiency can be met in any desired fashion. Further, the reactions taking place in the isomerization-craclcing zone entail a net reduction in the aromatic content of the stock as measured by comparing the aromatic content of the yfeed (including recycle) with that of the synthetic product portion of the efduent from the said reaction zone. Despite this reduction in aromatic content, however, the octane number of synthetic gasoline product is well above that of the corresponding feed fraction, and the product is reforrnable in good yield to a high octane gasoline product. Moreover, the yield of said product is extremely high, and normally ranges from about to 110 volume percent in terms of the volume of (liquid) fresh feed supplied to the isomeriZation-cracking zone. From this it is obvious that the process is an extremely efficient one entailing little degradation of the feed to less valuable, normally gaseous components.

When, contrary to the method described above, a naphtha feed stream of relatively wide boiling range (e.g., F. or more) is not divided into rela-tively narrowboiling fractions for separate processing, but is fed whole to the isomerization-cracking zone, it is not possible to recover more than a lsmall precentage of the synthetic product present in the effluent from said zone, i.e., that boiling below the initial boiling point of the feed-. The rest or the synthetic product boils in the same range as portions of the remaining, unconverted feed and thus cannot be separated therefrom. Accordingly, if all lthe feed is to be converted to synthetic product in such method of operation, it is necessary -to further convert alreadyformed synthetic product to still lower-boiling product fractions, and this represents an entirely uneconomic method.

As feed to be fractionated and catalytically converted in accordance with the process of the present invention, there can be employed any thermally cracked, catalytically cracked or straight run naphtha or gas oil fractions boiling in the range of from about to 850 F., preferably 175 to 650 F., and having a boiling range or spread, of at least 100 F. in addition to possessing the desired boiling point range characteristics, the feed stocks employed in lthis invention should be low in nitrogen content and preferably should contain less than 200 ppm. nitrogen, and more preferably less than 50 ppm. This specification is somewhat flexible since higher nitrogen contents can be tolerated with the higher boiling fractions than with the lighter portions of the feed. In the case of stocks which are not already sufficiently .low in nitrogen, acceptableV levels can be reached by pretreating the distillate (either before orafter fractionating the same in the manner described herein) with hydrogen in the presence of a suitable catalyst, for example cobaltand molybdenum-containing alumina catalyst at elevated temperatures and pressures. A particularly effective catalyst for this purpose is one wherein a coprecipitate molybdenaalumina material, as prepared in accordance with the disclosure of US. Patent No. 2,432,286 to Claussen et al., is combined with cobalt oxide, the final catalyst having a metals content of about 2% cobalt and 7% molybdenum. RepresentativeV processing conditions for removing nitrogen with this catalyst are 60G-800 F., SGO-2000 p.s.i.g., and 100G-6000 standard cubic feet of hydrogen per barrel of feed stock.

As noted above, the .feed stock employed should be fractionated into at least two and preferably into three or more fractions, each having an appreciable boiling range. *'Normally, this range or spread, will be of the order of 50 to 125 F., with the number of fractions employed being such that all portions of the given feed stock fall into one fraction or another. stocks boilirror in major portion within the range of from In the case ofy Y and similar materials.

Y operating conditionsl 3 about270 to 450 F., at least one fraction Vshould be chosen so as to contain the C3 aromatics and other components boiling somewhat below the C9 aromatics present V(a Ysuitable cut, for example, being one boiling from about 270 to 320 Fi about 10), while another fraction should be selected so as to contain the C9 aromatics together with other hydrocarbons present normally boiling between about 320 and 370 F., again-* -about 10. Additionally, one or more other Afractions' may be `obtained boiling above the latter fraction, as is also the case with the portion of the feed, if any, boiling below 270 F.

The catalyst employed in tie isomerization-cracking zone is one wherein a lmaterial having hydrogenatingdehydrogenating activity is deposited or otherwise disposed on an active cracking catalyst support. The cracking component may comprise any one or more ,of such acidic *materials as silica-alumina,y silica-magnesia, silica-alumina-zirconia composites, alumina-boris, fluorided composites, Vand the like, as well as various acid treated clays one or more of the various Groups V, VI, Vil and Jill metals, as well as from theioxides and suldes thereof, Valone or together with promoters or stabilizers that may have by themselves small catalytic effect, representative materials being oxides and suldes of molybdenum, tungsten, vanadium, chromium andthe like, as well as of metals such as iron, nickel, cobalt and'platinum. If desired, more than one hydrogenatingdehydrogenating component Vcan be present, and good results have been obtained with o Acatalysts containing ycomposites of two or more Vof the oxides of molybdenum, cobalhrnickel, chromium and zinc, .and with mixtures ofV said oxides with fluorine. The amount of the hydrogenating-dehydrogenating compouent present can be varied within relatively wide limits of from 0.5 to 30%, preferably 0.5 to 15%, based on the weight of the-entire catalyst.

required conversion levels, Vbut insulhcient to electsubstantial saturation of anytexcept highly substituted and polynuclear aromatic rings under the reaction conditions employed in the isomeriaztion-cracking zone.

Having selected or prepared the low nitrogen feed stock to` be processed, and split the same into the desired por# tions, the process of the present invention can be carried out in a continuous fashion by preheating the feed and added hydrogen supplied therewith to the extent necessary to bring thereaction in the isomerization-cracking zone on-streamk at the desired temperature about 400 Thereafter, while maintaining pressures and feed throughpu-t ratesiat the desired levels, the temperature in said -zone is controlled so as -to maintain (average) reaction temperatures therein at opti-mum levels. The normal f The foregoing remarks presupposed the employment of a' xed hed method of operation, and the same is preferred. However, the process of this invention can also be practiced using'a moving catalyst zbed or one of the fluidized type where catalyst regeneration is continuous andY on-stream periods can be indefinitely prolonged.A

In theicase of each feed fraction, the 'euent from the isomerizationcracking zonefis freed of a hydrogen-richV Vstream(normally recycled to said Vvzone) and of otherV f normally gaseous components, with the remainder then Y The hydrogenating-dehydrogenat-V ing components of the catalyst can be selected from any being'pass'ed to a fractionation zone for recovery of the desired (normally synthetic) product fraction or fractions. The remaining, higher-boiling material, in the case Y of each ethuent stream, can then be withdrawn-in whole or part from the system it" desired (for use as a component of various gasoline, jet or diesel fuels) though it preferably is recycled to the isomerization-cracking' zone for further conversion therein, along with added fresh feedV and hydrogen. The synthetic gasoline product fractions recovered from the various elluent streams can be combined, it desired, to produce a high octane gasoline of full boiling range, or theymay be separate-1y employed'in any desired fashion, usually as premium blending stocks. i.

A. process flow for a specic embodiment of this invention is illustrated in the appended drawing, and a detailed i description of a-'preferred manner of operationwill be discussed below in connection w1tn that drawing.

A'distillate feed stream (for example one having aV boiling point range of 230 to 380 F.) is passed through line lil into a fractionating column lll which'separates the naphtha into a light fraction boiling betweenabout 230 and 280 E. whichy is taken overhead through line 13,.

an intermediate Vfraction boiling between about 280 and 330 F. takenY through line 14, and the bottoms fraction boiling between 330 and 380 F. withdrawn through line l5. The fraction in line 13, along with hydrogen supplied through line i6, is heated to reaction temperature'by Y Within these limits, Vthetmount of said component present should be sufficient 'tol provide a reasonable catalyst ori-stream period at passage through a heat exchanger 17, the heated admixture so produced then being passed into reaction zone l Y where it is contacted with a clay-type cracking catalyst incorporating molybdenum oxide 4as the hydrogenation catalyst. The effluent from this reaction zone is passed via line i9 through a condenser 20 and into a gas-liquid separator 2l. Hydrogen, along' with a small percentage of other gaseous components, isV withdrawn from separa- Y Ytor 2l and returned through line 22 to the :hydrogen supply line 16, Whilel the liquid condensate in the Vseparator is withdrawn through line 23tand passed into distil-V Y lation zone' Z5 forrecovery of products. Normally gas-` eous materials are removed from zone 25 throughline 27,

while a gasoline fraction `boiling below the initial-boiling point (230 F.)-of the feed fraction inline i3 is recovered through line 28. The bottoms from zone 25 are recovered through line 29 and are'returned to line 13 Y for repeated catalytic conversion in zone 18.

The processing of the feed fractions in lines E4V and 15 is similarV to that of the fraction in line 13 as described Y above, said feeds passing, along with hydrogen supplied through lines 30 and 3l, through heat exchangers 32 and 33 to reaction zones 34 and 35 provided with catalyst of similar nature to that vpresent in'zone 18. The respective effluent streams from the latter reaction zones `are passed 'via lines 36 and 37 through condensers 3S and 39 to gas- Vliquid separators .40 and '41 from which hydrogenis taken overhead through lines iZYfand 43 and returned to the respective khydrogen supply ylines 30 and 31. liquid condensate streams from the separators are withdrawn throughlines 44 and 45 and passed into respective Vdistillation zones 46 `and 47, provided with lines 43 and Y 49 yfor release of normally gaseous materials. Fromzthe Zone 46 a gasoline fraction boiling below 280 R is taken through line 50 and joined with that from line 2S, while from zone 4 7 a gasoline stream boiling below 330 F. is taken through line 51 land also joined with that from vline 2S, the combined stream in line 52 representing the high quality product which, after suitable debutanizing and other steps, if required, is ready for use asfa premium gasoline havingarelatively wide'boiling range. The bot- 1 toms from theV zones 46 and47 aretaken through lines 53 and 54 andare recycled backV to lines i4 and 15, respectively. `f

In the flow diagram, vvariousrof Vthe required pumps, heat exchangers, valves, recycle-systems and other items lof flow control equipment have beenomitted inthe interest of simplicity andclarity of expression, the placing The i.

of such equipment being evident to those skilled in the art once the general procedure has been outlined. The process of the invention is illustrated by the following examples.

EXAMPLE 1 In this operation there was employed a naphtha having the following inspections:

A.P.1. gravity 48.1 Aniline point F 108 Volume percent aromatics 23 Volume percent paraifins plus naphthenes 76 F-l clear octane number 60.8

For the purpose of this example, a total of 8 volume percent constituting the fractions boiling below 180 F. and above 430 F. was removed from the above naphtha by rerunning. The remainder, boiling between 180 and 430 F., and having an aromatic content of 23% by volurne, was distilled with good fractionation into ve frac- Each of the foregoing fractions was then passed at a rate of 2 v./v./hr., at 800 F. and 1200 p.s.i.g., along with 6000 standard cubic feet of hydrogen per barrel of feed, through a catalyst comprised of molybdenum oxide (1% Mo) deposited on a silica-alumina cracking catalyst. Each conversion stream was then fractionated so as to recover the product fraction boiling below the initial boiling point of the particular naphtha cut involved, with the bottoms then being recycled back through the reaction zone, along with added feed and fresh as well as recycle hydrogen. Each of the lower boiling (gasoline) fractions thus recovered was then freed of C4 and lighter components, the yield and quality of each resulting, debutanized product portion produced being indicated below in Table I.

1 Product boiling below initial boiling point of feed fraction.

The blend obtained on combining the synthetic, debutanized products obtained from the above operation has an aromatic content of 20 volume percent which compares with an aromatic content of 23 volume percent in the'entire 180 to 430 F. naphtha feed. The linal yield of C-380" F. gasoline produced (all as synthetic product) lis equivalent to 70 volume percent, while the C4,- 380" F. yield is equivalent to 102.7 volume percent, both in terms of the volume of 180-430" F. naphtha feed. The overall F-l clear octane number of the C5-380 F. synthetic product is 88.2.

When, under the same conditions of temperature, pressure, throughput rate, hydrogen/feed ratio and catalyst, the whole M80-430 F.) fraction is processed in the 6 fashion described above it is found that the F-l clear octane rating of the whole eiiiuent, on a once through basis, is raised from a value of approximately 60 to 69.0. By fractionating this whole naphtha efliuent, it is not possible to recover more than about 14 volume percent of a debutanized synthetic gasoline product (having an F-l clear octane rating of 88), and this product is undesirable since its end point is low (about 180 F.). The yield of gasoline above 80 F-l clear octane number is below 25 volume percent. These yields are in contrast with that of to a gasoline of full boiling range and 88.2 F-l clear octane number, as obtained in the selective operation described above upon which the present invention is based.

EXAMPLE 2 A cycle oil boiling in the range 400 to 850 F. is separated into two fractions, a light fraction boiling from 400 to 525 F. and a heavy fraction boiling from 525 to 850 F. The light fraction is separately hydrocracked in a rst conversion zone over a catalyst comprising nickel sulfide on silica-alumina at a space velocity of 2 v./v./hr., a temperature of 700 to 800 F., and a pressure of 1300 p.s.i.g., in the presence of 6000 standard cubic feet of hydrogen per barrel of feed. The heavy fraction also is separately hydrocracked, in a second conversion zone, over the same type of catalyst and under the same conditions as used in the first conversion zone. The yields from each of the two hydrocracking Zones are as follows:

Heavy Light Total Feed Feed Weight Percent Volume Percent The efuent from the first conversion Zore is separated into a gasoline product fraction and a fraction boiling in the range 400 to 525 F. which is recycled to the rst conversion zone.

The eliiuent from the second conversion zone is separated into a gasoline product fraction, a middle distillate fraction boiling from 400 to 525 F., and a fraction boiling above 525 F. which is recycled to the second conversion zone.

The middle distillate fraction recovered as a product from the second conversion zone is highly useful for jet fuel purposes, particularly because of its low freeze point. The first crystal point of the product is below -70 F.

If the entire original 400 to 850 F. feed stock is not iseparated as aforesaid into a lighter and a heavier portion which are separately hydrocracked, but rather is hydrocracked in toto in a single conversion Zone, the 400 to 525 F. middle distillate product has a considerably higher freezing point, because it contains unconverted light cycle oil which has a first crystal point of about 0 F. Le., the 400 to 525 F. material recycled to the first conversion zone above has a first crystal point of about 0 F. and if even a small amount of this material is mixed with the 400 to 525 F. middle distillate product from the second conversion zone, the freeze vpoint of that product would be raised in an intolerable amount, perhaps as much as 50 F.

We claim: l. A process for upgrading a petroleum distillate boiling in the range of from Vabout to 850 F; and having range; separately passing each of said'feed fractions, along with at leaStj1500's.c.f. of hydrogen per barrel thereof,

. in contact with a catalyst comprising at least one hydrogenatingdehydrogenating component and an active cracking catalyst support in an isomerization-cracking zone ing from the effluent fromsaid yisomerization-cracking zone, as a product of the conversion reaction occurringV therein, a product fraction boilingV essentiallyV below the initialrboiling point of the particular feed fraction employed; said conversion reaction being accompanied by a net consumption of -from about 800 to 3000 s.c.f. of

, hydrogen per barrel of feed converted to synthetic product in said'zone.

2. The process of claim l, wherein the portions of said eiuent boiling above the end point of said recovered product fraction are recycled, at least in part, to said isomerization-cracking zone.

V3. A process for upgradinga petroleum naphtha boiling in a range of from about 175 to 850 F. and having a boiling range of at least 100 F., which comprises fractionally distilling said napht'ha to separate the same into at least two feed fractions, each having an appreciable boiling range, separately contacting each of said fractions, in the presence of at least 1500 s.c.f.of hydrogen per barrel thereof, at pressures above 600 psig, and at temperatures between about 400 and 900 F., with a catalyst comprised of at least one hydrogenatingdehydrogenating component and an active cracking catavlystsupport;separating the hydrogen and other normally Vao of each of said streams into a bottoms fraction and an overhead product fraction boiling essentially below the initial boiling point of the feed fraction involved; said process being characterized by a net consumption of from about 800 to 3000 s.c.f. of hydrogen for each barrel of feed converted to synthetic product by said contact with the catalyst.

4. The process of claim 3, wherein a hydrogen-rich gas separated. from said reaction product streams is refV cycled back through the catalyst along with added hydro-V gen, and wherein said bottoms fraction is also recycled backthrough the catalyst along with added quantities of the corresponding feed fraction from which said bottoms fraction derives.

5. The process of claim 4, wherein said feed fractions each have a boiling'range of from about 50 to 150 F.

6. The process of claim 4, wherein said petroleum naphtha is one which boils in major portions within the range of from about 270 to 450 F., with a boiling range of at least F., Yand Which is fractionated into at least two fractions, one of which contains C8 aromatics and boils in the range of from about 270y to 320 F., i about 10, and another of which contains C9 aromatics and Vboils in the range of from about 320 to 370 F., :t: about 10F.

neferens cited in the nie of this' patent UNITED STATES PATENTS 2,241,430 vSnow May 13, 1941 2,773,809 Haensel et al. Dec. l1, 17956 2,866,745 Heinemann Dec. 30, 1958 2,937,132 Voorhies May 17, 1960 2,944,006v Scott uly ,5,V 1960 V2,)44,959 Kline et al. uly'liZ, 1960 

1. A PROCESS FOR UPGRADING A PETEROLEUM DISTILLATE BOILING IN THE RANGE OF FROM ABOUT 175 TO 850*F. AND HAVING A BOILING RANGE OF AT LEAST 100*F., WHICH COMPRISES FRACTIONATING SAID DISTILLATE TO SEPARATE THE SAME INTO AT LEAST TWO FEED FRACTIONS EACH AHVING AN APPRECIABLE BOILING RANGE; SEPARATELY PASSING EACH OF SAID FEED FRACTIONS, ALONG WITH AT LEAST 1500 S.C.F. OF HYDROGEN PER BARREL THEREOF, IN CONTACT WITH A CATALYST COMPRISING AT LEAST ONE HYDROGENATING-DEHYDROGENATING COMPONENT AND AN ACTIVE CRACKING CATALYST SUPPORT IN AN ISOMERIZATION-CRACKIN ZONE FOR A GIVEN ON-STREAM PERIOD, SAID ZONE BEING MAINTAINED AT A PRESSURE OF AT LEAST 600 P.S.I.G. AND AT AN AVERAGE TEMPERATURE BETWEEN ABOUT 400 AND 900*F., AND RECOVERING FROM THE EFFLUENT FROM SAID IXOMERIZATION-CRACKING ZONE, AS A PRODUCT OF THE CONVERSION REACTION OCCURRING THEREIN, A PRODUCT FRACTION BOILING ESSENTIALLY BELOW THE INITIAL BOILING POINT OF THE PARTICULAR FEED FRACTION EMPLOYED; SAID CONVERSION REACTION BEING ACCOMPANIED BY A NET CONSUMPTION OF FROM ABOUT 800 TO 3000 S.C.F. OF HYDROGEN PER BARRLE OF FEED CONVERTED TO SYNTHETIC PRODCUT IN SAID ZONE. 